Production of methane-rich gas

ABSTRACT

Methane-rich gas is produced by reacting gas mixtures comprising CO and at least one member selected from the group consisting of H 2  and H 2  O, such as synthesis gas, in a reactor containing an improved unsupported catalyst comprising an alkali-metal promoted partially reduced mixture of at least one nickel uranate and at least one oxide of nickel. The weight percent uranium present in the activated catalyst, based on total weight of catalyst composition, is in the range of over 50 to about 90.

BACKGROUND OF THE INVENTION

1. Field of the Invention

This invention relates to an improved process and catalyst for theproduction of methane-rich gas. More specifically, the present inventionrelates to the production of a product gas stream comprising at least 25vol. % methane (dry basis) by reacting a feed gas stream, such as thatproduced by the partial oxidation of a liquid hydrocarbonaceous or solidcarbonaceous fuel, in a reactor containing the improved catalyst. Theproduct gas may be burned as a fuel without polluting the environment.

2. Description of the Prior Art

Synthesis gas containing methane may be made by the partial oxidation ofa hydrocarbonaceous fuel using comparatively high steam to fuel weightratios and no subsequent catalytic methanation step, as described incoassigned U.S. Pat. No. 3,688,438. In coassigned U.S. Pat. No.3,890,113, a gaseous stream comprising H₂ O and CO is produced by thepartial oxidation of a hydrocarbonaceous fuel and is subjected to thewater-gas shift reaction to produce a gaseous stream rich in H₂ and CO₂.This gas mixture is then subjected to conventional catalytic methanationafter the mole ratio H₂ /CO₂ is adjusted to about 4 to 10. In coassignedU.S. Pat. No. 3,888,043, a large amount of nitrogen diluent is containedin the effluent gas stream from the partial oxidation generator. Duringthe two conventional methanation steps with an intervening water-gasshift conversion step, the large amount of nitrogen diluent in thereacting gas stream helps to control the normally vigorous exothermicreactions which are going on.

In U.S. Pat. No. 4,032,556, the methane content of a gas produced by thehydrogenation of a liquid hydrocarbon with a hydrogenating gas at hightemperatures is increased by contacting the gas stream with a supportednickel-urania hydrogenation catalyst comprising a maximum of 10 wt. %uranium. U.S. Pat. No. 3,993,459 pertains to a catalyst for convertinghigher hydrocarbons into gas mixtures containing carbon monoxide,methane and/or hydrogen in which the active components are oxides of themetals lanthanum, cobalt, nickel, uranium, cerium and thorium on anoxide substrate. The catalyst comprises from about 0.1 to 8 percent byweight of uranium. In U.S. Pat. No. 3,847,836 liquid hydrocarbons e.g.naphtha are steam reformed using a catalyst comprising nickel, and/ornickel oxide, together with a relatively smaller amount of uranium oxidesupported on a carrier.

SUMMARY

This is a process employing an improved catalyst for producing amethane-rich gas comprising at least 25 vol. % methane (dry basis) froma CO-containing gas mixture. The catalyst is operable over a widetemperature range. The feed gas stream comprises CO and at least onemember selected from the group consisting of H₂ and H₂ O in which themole ratio of either H₂ /CO or H₂ O/CO ranges from about 0 to 10.0, suchas 0 to 5, and the remaining ratio is in the range of about 0.3 to 10,such as 0.3 to 5. Preferably, the raw feed gas may be obtained from thepartial oxidation of liquid hydrocarbonaceous or solid carbonaceousfuels. In such case in addition to CO and H₂ and/or H₂ O, the feed gasmay also comprise at least one member selected from the group consistingof CO₂, CH₄, N₂ and Ar. The high activity of the subject catalystpermits removal of trace amounts of CO from a gas mixture.

The feed gas is reacted in a fixed bed reactor containing an improvedunsupported catalyst consisting essentially of an alkali-metal promotedpartially reduced mixture of at least one nickel uranate and at leastone oxide of nickel. The weight percent uranium present in the activatedcatalyst, based on the total weight of the catalyst composition, is inthe range of over 50 to about 90, such as about 60-80; and the weightratio U/Ni is in the range of about 7-1, such as about 5-2, say 3.Preferably, the uranium in the catalyst composition comprises at least99.7 wt. % ²³⁸ U isotope. The catalyst has a low reaction initiationtemperature, i.e. about 375° F. and is highly active up to about 1500°F. Water-gas shift and methanation reactions may take placesimultaneously in the reactor. While the dry methane-rich product gascomprises at least 25 vol. % methane, it may range up to about 97 vol. %methane on a single pass basis.

BRIEF DESCRIPTION OF THE DRAWING

The invention will be better understood by reference to the drawing,which illustrates one embodiment of the disclosed process.

DESCRIPTION OF THE INVENTION

The present invention pertains to an improved continuous process for theproduction of a gaseous stream comprising 25 to about 97 vol. % CH₄ (drybasis) or more by the catalytic reaction of CO and at least one memberof the group consisting of H₂ and H₂ O. An improved, unsupportedcatalyst consisting essentially of an alkali-metal promoted, partiallyreduced mixture of at least one nickel uranate and at least one oxide ofnickel is employed as the catalyst in the subject process.

A particular advantage of the subject process is that the gaseousfeedstock may be produced from readily available, comparatively lowcost, hydrocarbonaceous and carbonaceous materials e.g., liquid andsolid fuels which may contain a comparatively high content of ash andsulfur. When necessary, the feed gas is purified prior to beingintroduced into the catalytic reactor. The product gas has a heatingvalue in the range of about 400 to 1000 British Thermal Units perstandard cubic foot (dry, CO₂ -free basis), depending upon the methanecontent. The product gas may be used as a substitute for natural gas orin organic chemical synthesis when the methane content is greater than95 mole %.

The feed gas being catalytically reacted in the subject process may varyin composition from mixtures of pure CO and H₂ O to mixtures comprisinga trace of CO in an H₂ stream. Thus, the feed gas stream for producing aCH₄ -rich gas comprises carbon monoxide and at least one member selectedfrom the group consisting of H₂ and H₂ O in which the mole ratio ofeither H₂ /CO or H₂ O/CO ranges from about 0 to 10.0, such as 0 l to 5,and the remaining ratio is in the range of about 0.3 to 10, such as 0.3to 5.

The feed gas may be produced by the partial oxidation of ahydrocarbonaceous or carbonaceous fuel followed by gas purification. Insuch case in addition to CO and H₂ and/or H₂ O, the feed gas stream mayalso contain at least one member selected from the group consisting ofCO₂, CH₄, N₂ and Ar. It was unexpectedly found that carbon may beprevented from depositing on the catalyst in the reactor by maintainingthe feed gas stream with a minimum atomic ratio H/C of 2. The source ofthe hydrogen may be H₂ or H₂ O.

Any conventional source of CO in admixture with H₂ and/or H₂ O may beemployed for the feed stream to the catalytic reactor. For example,catalytic steam reforming and particularly the partial oxidation processare suitable. The partial oxidation of a hydrocarbonaceous fuel i.e.liquid hydrocarbons or solid carbonaceous fuel such as coal with afree-oxygen containing gas with or without a temperature moderatorproduces mixtures of H₂ and CO in which the mole ratio H₂ /CO may varyin the range of about 0.30 to 5 depending on the source of carbon used.While the above customary convention has been adopted herein forexpressing ranges of ratios, the aforesaid range of ratios may also beexpressed in the following manner: 0.30/1 to 5/1. In general, westerncoals and lignites will yield gas mixtures in which the mole ratio H₂/CO may vary in the range of about 0.4 to 0.6. In another embodiment,the high activity of the subject catalyst permits the removal of traceamounts i.e. up to 2 mole % of CO from the gas mixture.

Conventional methods of making synthetic natural gas (SNG) require amole ratio H₂ /CO in the range of about 1 to 3 as shown by the followingequations:

    CO+3H.sub.2 →CH.sub.4 +H.sub.2 O                    (1)

    2CO+2H.sub.2 →CH.sub.4 +CO.sub.2                    (2)

In order to realize such a ratio, in many prior art processes a portionof the gasification product containing CO is first catalytically reactedwith H₂ O in a separate reactor over a separate water-gas shift catalystto produce H₂ and CO₂. The CO₂ is then removed, and the H₂ is added tothe original stream to produce the desired ratio of H₂ /CO. The gasmixture is then reacted over a separate methanation catalyst in aseparate reactor to produce CH₄.

In contrast, it has been found that the subject improved catalystunexpectedly acts simultaneously as a water-gas shift catalyst and as amethanation catalyst. Advantageously by this means, only a single bed ofthe subject catalyst is required. This permits significant savings inthe cost of catalyst and equipment. Further, wasteful intermediate gascooling and purification steps may be eliminated. Process control isless critical as the mole ratios H₂ /CO and/or H₂ O/CO in the feed gasmay vary over a wide range.

A wide variety of hydrocarbonaceous fuels is suitable as feedstock forthe partial oxidation process, either alone or in combination with eachother or with particulate carbon. The hydrocarbonaceous feeds include bydefinition fossil fuels such as various liquid hydrocarbon fuelsincluding petroleum distillates and residua, naphtha, asphalt, gas oil,residual fuel, reduced crude, fuel oil, whole crude, coal tar, coalderived oil, shale oil, tar sand oil and mixtures thereof. Suitableliquid hydrocarbon fuel feeds as used herein are by definition liquidhydrocarbonaceous fuel feeds that have a gravity in degrees API in therange of about -20 to 100.

Pumpable slurries of solid carbonaceous fuels, e.g., lignite, bituminousand anthracite coals, coal char, particulate carbon, petroleum coke, andmixtures thereof in water or in said liquid hydrocarbon fuels areincluded herewith as within the scope of the definition forhydrocarbonaceous fuel feeds.

Further, included also by definition as a hydrocarbonaceous fuel areliquid oxygenated hydrocarbonaceous materials i.e., liquid hydrocarbonmaterials containing combined oxygen, including alcohols, ketones,aldehydes, organic acids, esters, ethers, and mixtures thereof. Further,a liquid oxygenated hydrocarbonaceous material may be in admixture withone of said liquid petroleum materials.

The term free-oxygen containing gas or gaseous oxidant as used herein isintended to mean a gas selected from the group consisting of air,oxygen-enriched air (22 mole percent O₂ and higher), and preferablysubstantially pure oxygen (95 mole percent O₂ and higher). The amount ofnitrogen in the product gas may be substantially reduced or eliminatedby using substantially pure oxygen. The ratio of free-oxygen in thegaseous oxidant to carbon in the feedstock (O/C, atom/atom) is in therange of about 0.6 to 1.5, suitably about 0.7 to 1.2 and preferablybelow 1.0.

H₂ O is preferably introduced into the reaction zone to help control thereaction temperature, to act as a dispersant of the hydrocarbonaceousfuel fed to the reaction zone, and to serve as a reactant to increasethe relative amount of hydrogen produced. About 0.15 to 5.0 pounds of H₂O are introduced per pound of hydrocarbonaceous fuel. Other suitabletemperature moderators include CO₂ -rich gas, a cooled portion ofeffluent gas from the gas generator, cooled off-gas from an integratedore-reduction zone, nitrogen, and mixtures thereof.

The subject improved catalyst comprises an alkali-metal promotedpartially reduced mixture of at least one nickel uranate selected fromthe group consisting of NiUO₄ and NiU₃ O₁₀, and at least one oxide ofnickel selected from the group consisting of NiO and Ni₃ O₄. The totalamount of uranium present in the partially reduced catalyst, based onthe total weight of the partially reduced i.e. activated catalystcomposition, is in the range of over 50 to about 90, say about 60 to 80weight percent. The weight ratio U/Ni of the activated catalyst is inthe range of about 7-1, such as 5-2, say 3.0.

Preferably, the uranium in the present catalyst composition comprises atleast 99.7 weight percent ²³⁸ U isotope. Advantageously, the tailingsproduced by conventional ²³⁵ U enrichment processes may be easilyconverted into U₃ O₈ in which at least 99.7 wt. % of the uranium ispresent as ²³⁸ U isotope. Thus, tailings which were once a burdensomewaste material may now be used as a low cost raw material for theproduction of the catalyst of the subject invention.

The expression "alkali-metal" promoter includes at least one memberselected from the group consisting of potassium, sodium, and cesium.Typical examples of alkali metal compounds that may be used in theproduction of the subject catalyst composition to supply said alkalimetals are the respective alkali metal oxides or the salts ofoxygen-containing acids such as carbonates, bicarbonates, nitrates,oxalates and acetates, and/or hydroxides of said alkali metals whichyield the oxides at elevated temperatures. The preferred alkali metal inthe activated catalyst is potassium in the form of the compound K₂ O.The alkali metal compound content of the catalyst may range from about0.01 to 2 weight percent, say about 0.5-1 weight percent based on thetotal weight of the catalyst.

Most conventional catalysts employ a suitable support material such asaluminum oxide, silicon oxide, or an oxide of magnesium, calcium, orbarium. In contrast, the subject improved bulk catalyst is unsupported.A completely inactive catalyst resulted when the ingredients of thesubject catalyst were incorporated on said supporting materials.

The subject catalysts are prepared, primarily, by thermally decomposingnickel nitrate hexahydrate and an alkali metal compound in the presenceof at least one oxide of uranium selected from the group consisting ofU₃ O₈, UO₃, and UO₂. A mixture consisting essentially of at least onenickel oxide, and at least one nickel uranate, and an alkali metal oxideis thereby produced, such as the following mixtures of compounds inweight percent:

Formula A--NiO 1.0 to 30, Ni₃ O₄ trace to 30, NiUO₄ 70 to 98, and K₂ O0.01 to 2.0.

Formula B--NiO 1.0 to 30, NiU₃ O₁₀ 70 to 98, and K₂ O 0.01 to 2.0.

Formula C--NiO 1.0 to 30, NiUO₄ 30 to 68, NiU₃ O₁₀ 30 to 68, and K₂ O0.01 to 2.0.

It is necessary to partially reduce the aforesaid mixtures by treatmentwith pure hydrogen at high temperatures and pressures in order toproduce the activated catalyst of the subject invention. After saidtreatment with hydrogen, at least 30 up to about 80, say 40 to 60 wt. %of the nickel present in the activated catalyst is in the metallicstate. From about 85 to 100, such as at least 90 wt. % of the remainderof the nickel present in the activated catalyst may be in the form of atleast one nickel uranate selected from the group consisting of NiUO₄ andNiU₃ O₁₀. The balance of nickel, if any, may be present in the activatedcatalyst in the form of at least one oxide of nickel selected from thegroup consisting of NiO and Ni₃ O₄. From about 10 to 100, such as about20 to 80, say at least 90 wt. % of the uranium in the activated catalystmay be present in the form of at least one nickel uranate selected fromthe group consisting of NiUO₄ and NiU₃ O₁₀. The balance of uranium, ifany, may be present in the activated catalyst in the form of at leastone oxide of uranium selected from the group consisting of U₃ O₈, UO₃and UO₂.

In the preparation of the subject catalyst, about 29.1 g to 2,910 g suchas 0.1 to 10 moles of nickel nitrate hexahydrate (Ni(NO₃)₂.6H₂ O) areheated in a stainless steel vessel to a temperature in the range ofabout 130° to 200° F., such as 134° to 150° F. and melted. The materialis stirred as it is heated and 0.05 to 18.0 g, such as 0.0036 to 0.1333moles of an alkali metal compound having the formula M-A wherein M is analkali metal selected from the group consisting of Na, Ce, and K, and Ais a member selected from the group consisting of CO₃, HCO₃, NO₃, OH,oxalate and acetate, are added to the melt. The mixture is continuallystirred at the melting temperature and 84 g to 8420 g, such as 0.1 to 10moles, of at least one uranium oxide selected from the group consistingof U₃ O₈, UO₃, and UO₂ are gradually added over a period of about 15 to25 minutes.

After emission of nitrogen oxides and frothing ceases, the productthickens to a brown mass having a rubbery consistency. Mixing iscontinued from about 20 to 100 minutes. The mass is then cut up intosmall chunks and then heated for 16 to 24 hours at a temperature in therange of about 50° to 200° F. The chunks are then dried for 1 hour in aforced air oven at a temperature in the range of about 160° to 350° F.,and then heated in a muffle furnace to a temperature of about 400° F.The temperature is raised 100° F. per hour until a temperature in therange of about 675° to 725° F. is reached and maintained for about 2.5to 3.5 hours.

At this time the material is in the form of a powder whose particle sizemay be further reduced, by crushing, to about 100 to 1000 microns, suchas 150 to 600 microns, say 300 microns. Alternately, the powder may bemixed with a binder, such as stearic acid and molded into pellets havinga diameter in the range of about 1/16" to 1/4", say 5/32". The pelletsare heated in a muffle furnace under a slow stream of nitrogen to atemperature in the range of about 550°-650° F. for 2-3 hours, and thenat 750°-850° F. for an additional 3-4 hours.

The catalyst is activated by being partially reduced with pure hydrogen.1500 to 6500 standard cubic centimeters per minute of H₂ per 100 cc ofcatalyst are passed through the catalyst in a reactor for 10-15 hours ata temperature in the range of about 850°-900° F. and 200-250 psigpressure. The pressure is then increased to about 310 to 350 psig andthe hydrogen treatment is continued for an additional 6-8 hours at thesame conditions of flow rate and temperature.

Conventional fixed bed, ebullient bed, or fluidized bed reactors may beused for converting the feed gas into the methane-rich gas stream.Preferably, a fixed bed reactor is employed. Since the reaction ishighly exothermic, temperature control may be effected by any of thefollowing techniques: distribution of the feed gas throughout the fixedbed reactor by means of separate inlet points, imbedding tubular coolersin the catalyst beds and producing steam which may be used elsewhere inthe process, or cooling the effluent gas between beds, with simultaneoussteam generation.

Advantageously, no recycle is required with the subject highly activecatalyst. However, optionally about 1-25, say about 2-5, volumes of theproduct gas may be mixed with each volume of fresh feed gas. Thetemperature of the inlet feed gas stream to the reactor may be in therange of about 375°-725° F., such as about 395°-600° F. The catalyst bedtemperatures are not critical and can be allowed to rise as much as 775°F. and still produce near theoretical yields of CH₄. The catalyst isactive at the lower range and will start to make CH₄ and CO₂. The bedtemperature can climb rapidly to about 1500° F. and still produce neartheoretical amounts of CH₄.

Space velocities (standard volumes of dry gas per volume of catalyst perhour-vol./vol./hr) may be in the range of about 350-10,000, such as800-6000. The pressure in the reactor is in the range of about 1-300atmospheres, such as about 10-150 atmospheres. The reaction time is inthe range of about 1-100, such as 3-60 seconds. The catalyst has a longlife without appreciable deterioration. Carbon deposition may beresisted by maintaining an atomic ratio in the feed gas of H/C of atleast 2. The feed gas stream to the reactor may be preheated to theproper inlet temperature by indirect heat exchange with at least aportion of the effluent gas stream leaving the reactor at a temperaturefor example in the range of 600° to 1500° F. If necessary, the catalystmay be regenerated and most of its activity restored by hydrogentreatment.

With feedstreams comprising CO in admixture with H₂ and/or H₂ O, the CH₄-rich gas stream leaving the catalytic reactor comprises on the drybasis at least 25 mole % to 97 or more mole % CH₄. Such clean gas energyhas a heating value of about 400 to 1000 BTU/SCF. Using conventionalmethods the effluent gas from the reactor may be optionally dried andpurified by removing unwanted gaseous materials i.e. CO₂, H₂ S, COS, andN₂.

DESCRIPTION OF THE DRAWING

In one embodiment of the previously described process, as shown in theaccompanying schematic drawing, the feed gas to the CO catalytic reactor1 is originally produced in the reaction zone of free-flow unpackednoncatalytic partial oxidation synthesis generator 2 at an autogenoustemperature in the range of about 1700° to 3000° F. and a pressure inthe range of about 1 to 300 atmospheres. Note that, advantageously,there is no separate water-gas shift conversion step in the subjectprocess.

A hydrocarbonaceous feed in line 3 is passed through line 4 with orwithout a temperature moderator from line 5, valve 6, line 7 and thendown through annular passage 8 of burner 9 located in upper centralinlet 10 of vertical free-flow refractory lined noncatalytic gasgenerator 2. Simultaneously, a stream of free-oxygen containing gas inline 15 is passed down through central conduit 16 of burner 9.

Preheating of the reactants is optional but generally desirable. Forexample, a hydrocarbon oil and steam may be preheated to a temperaturein the range of about 100° to 800° F. and the oxygen may be preheated toa temperature in the range of about 100° to 1000° F.

The downflowing feedstreams impinge and the partial oxidation reactiontakes place in reaction zone 17. The raw effluent gas stream leaving thepartial oxidation gas generator through bottom outlet 18 passes throughinsulated unit 30 and may have the following composition in mole %: H₂20-70; CO 15-60; CO₂ 3-30; H₂ O 5-15; CH₄ nil-20; N₂ nil-60; H₂ S 0-5.0;COS 0-0.2; and Ar 0-2.0.

Unreacted particulate carbon (on the basis of carbon in the feed byweight) entrained in the effluent gas stream comprises about 0.2 to 20weight percent from liquid feeds but is usually negligible from gaseoushydrocarbon feeds. At least a portion of any entrained particulate solidmatter or molten or carbonaceous slag in the hydrocarbonaceous orcarbonaceous fuel may be separated from the effluent gas stream leavingthe gas generator in a suitable gas-solids separating zone withoutlowering the temperature or pressure of the gas stream. For example,particulate matter and slag, if any, may separate out in gas diversionand solids separation chamber 19 and pass through passage 20 into slagchamber 21. Slag chamber 21 may be connected in axial alignment with thefree-flow gas generator 2. By this means, ash and other solids or moltenslag in the gas stream discharging from the lower part of reactionchamber may drop directly into a pool of water contained in the bottomof the slag chamber. The separated material may be periodically removedthrough line 22, valve 23, and line 24.

The hot effluent gas stream leaving gas diversion and solids separationchamber 19 is passed through insulated line 30 and into gas cooler 31where it may be cooled to a temperature in the range of about 400° to800° F., say about 450° to 650° F. by indirect heat exchange with water.Boiler feed water enters gas cooler 31 by way of line 32 and inlet 33.The water is converted into steam which leaves by way of outlet 34 andline 35. By-product steam is thereby produced for use elsewhere in thesystem and/or for export. Alternatively, all of the hot raw effluent gassteam leaving the gas generator may be cooled by direct quenching inwater.

The partially cooled stream of effluent gas leaving gas cooler 31through outlet 36 and line 37, or the effluent gas stream leaving aquench tank not shown, is passed into a gas cleaning zone where anyremaining entrained solids may be removed. Any conventional procedurefor removing suspended solids from the gas stream may be employed. Forexample, the effluent gas stream may be passed through a scrubbingcolumn in direct contact and counterflow with a scrubbing fluid selectedfrom the group consisting of liquid hydrocarbon, dilute mixtures ofparticulates carbon and scrubbing fluid, or water. A slurry ofparticulate carbon and scrubbing fluid may be then removed from thebottom of the column and sent to a carbon separation or concentrationzone. Carbon concentration may be effected by any suitable conventionalmeans e.g. filtration, centrifuge, gravity settling, or by liquidhydrocarbon extraction such as the process described in coassigned U.S.Pat. No. 4,205,963. In the scheme shown in the drawing, the effluent gasstream is scrubbed with water in gas scrubber 38. The water entersthrough line 39 and the dispersion of solids and water leaves throughline 40 and is sent to a conventional facility not shown for therecovery of water and carbon. The gas stream in line 45 may be thenpassed directly through lines 58-61 and 57 into catalytic reactor 1.Optionally, recycle gas from line 71 may be mixed with the cooled andcleaned gas stream from line 60.

When necessary, at least a portion of the cooled and cleaned gas streamin line 54 may be treated to remove gaseous impurities i.e. acid-gas.The purified and any unpurified portions of the gas stream are thencombined in line 57 and fed to the reactor. In said embodiment, at leasta portion i.e. 50-100 volume percent of the clean gas stream leaving gasscrubbers 38 at a temperature in the range of about 375°-725° F. ispassed through lines 45, 46 and cooled below the dew point in gas cooler47. Water and normally liquid hydrocarbons, if any, may be therebycondensed out. The mixture of gas and liquids is passed through line 48and separated in gas-liquids separator 49. The liquids are removed fromseparator 49 by way of line 50 at the bottom. The substantially dry gasstream passes overhead through line 51 into the gas purification section52 at a temperature in the range of about 100°-175° F. Any suitableconventional process may be used for purifying the process gas stream ingas purification section 52. Typical gas purification processes mayinvolve refrigeration and physical or chemical absorption with asolvent, such as methanol, n-methyl-pyrrolidone, triethanolamine,propylene carbonate or alternatively with hot potassium carbonate. Forexample, the gas stream may be washed with cold methanol and the totalsulfur, H₂ S plus COS, may be reduced to less than 0.1 ppm. All of theCO₂ may be then removed to produce a purified process gas streamcontaining less than 5 ppm CO₂. The solvent is regenerated and recycledto the absorption column for reuse. Thus, rich solvent absorbent isremoved from gas purification section 52 by way of line 42, regenerated,and returned as lean solvent absorbent by way of line 43. One or more ofthe gases separated from the solvent absorbent during regeneration, andpreferably CO₂, may be recycled to the partial oxidation gas generatoras at least a portion of the temperature moderator. When necessary,final clean-up may be accomplished by passing the process gas streamthrough iron oxide, zinc oxide, or activated carbon to remove residualtraces of H₂ S or organic sulfide. Thus, at least one of the followinggaseous impurities when present in the process gas stream may be removedin gas purification section 52: CO₂, H₂ S, COS, Ar, N₂, and H₂ O.Optionally, a portion of the CO₂ and/or N.sub. 2 may be allowed toremain in the process gas stream to impart improved temperature controlwithin catalytic reactor 1. Optionally, at least a portion of the H₂ Omay be left in the process gas stream for the water-gas shift reaction.

A cleaned and purified gas stream leaves gas purification section 52through line 53 at a temperature in the range of about 70°-150° F. Thisgas stream is heated in heat exchanger 54 by indirect heat exchange withthe hot product gas leaving reactor 1 by way of line 55 at a temperaturein the range of about 600°-1500° F. The heated gas stream leaving heatexchanger 54 passes through lines 56, 57 and enters catalytic reactor 1at a temperature in the range of about 375°-725° F. The remainingportion, if any, of the cleaned gas stream leaving overheat in line 45of gas scrubber 38 and bypassing the gas drying and purificationsections by way of line 58, valve 59, and lines 60 and 61 is mixed inline 57 with the preheated cleaned, dried, and purified gas stream fromline 56. All or a portion of the partially cooled product gas streamleaving heat exchanger 54 by way of line 80 is passed through line 62and cooled below the dew point in gas cooler 63. Water and any traces ofnormally liquid hydrocarbons condense out. The mixed phase mixturepasses through line 64 into gas-liquids separator 65 where separationtakes place. Liquids are removed intermittently from separator 65 by wayof line 66 at the bottom. The methane-rich product gas leaves overheadfrom separator 65 by way of lines 67 and 68. A recycled stream, if any,is passed through line 69, valve 70, lines 71, 61, and mixed in line 57with the cleaned dried and purified gas stream from line 56. Thatportion of the partially cooled gas stream from line 80, if any, whichis not dewatered, is passed through line 81, valve 82, and line 83.

In one embodiment, the feed gas stream in line 57 contains H₂ O. H₂ Omay be added separately in the form of water or steam or alternatively,at least a portion of the H₂ O in the gas stream leaving gas scrubber 38may be allowed to remain. The mole of either H₂ /CO or H₂ O/CO in thefeed gas mixture in line 57 is in the range of 0 to 10, say 0 to 5, andthe remaining ratio is in the range of about 0.3 to 10, say 0.3 to 5.Reactor 1 is a fixed bed reactor containing the subject unsupportedcatalyst comprising a alkali-metal promoted partially reduced mixture ofat least one nickel uranate and at least one oxide of nickel. Thepressure in reactor 1 is substantially the same as that in partialoxidation gas generator 2 i.e. 1-300 atmospheres, say 10-200atmospheres, less the ordinary pressure drop in the lines.

The dry product gas in line 68 may have the following composition inmole %: CH₄ 25-98, CO₂ 1-75, H₂ nil-40, CO nil-10, N₂ nil-60, and Arnil-2.0. Any or all of the gaseous impurities may be removed byconventional gas purification methods not shown to produce clean fuelgas or methane. For example, CO₂ may be separated and at least a portionrecycled to synthesis gas generator 2 as at least a portion of thetemperature moderator.

EXAMPLES

The following examples are offered as a better understanding of thepresent invention, but the invention is not to be construed as limitedthereto.

EXAMPLE I

Example I will illustrate a preferred method for preparing a catalystcomposition which is activated in Example II according to thisinvention.

491 gms (1.69 moles) of nickel nitrate hexahydrate (Ni(NO₃)₂.6H₂ O) weremelted and mixed with 2.0 gms (0.0143 moles) of potassium carbonate (K₂CO₃). The mixture was maintained at the melting temperature andcontinuously stirred while 350 gms (0.417 moles) of triuranium octoxidepowder (U₃ O₈) were added gradually over a twenty minute period. Theproduct thickened to a brown mass having a rubbery consistency. The masswas cut into small chunks and allowed to remain on a steam plate forabout 20 hours. After being dried for 1 hour in a forced air oven at atemperature of about 250° F., the chunks were heated in a muffle furnaceto a temperature of 400° F. The temperature was then raised 100° F. perhour until 700° F. was reached and then maintained for 3 hours. By thattime the chunks had become a fine powder. The powder was further crushedto allow total passage through a 50 mesh screen (297 microns). Thepowder was then mixed with 10 grams of stearic acid powder and pelletedinto 5/32 inch diameter pellets. The pellets were heated in a mufflefurnace under a slow stream of N₂ for 2 hours, and then for anadditional 3 hours at a temperature of 800° F. The final catalyst wasdesignated Catalyst A, and had the following approximate chemicalcomposition in weight percent: NiO 6.5, Ni₃ O₄ trace, NiUO₄ 93.1, and K₂O 0.35. Catalyst A was activated in the manner described in Example II,prior to being used in Examples II to IV. In one embodiment, the uraniumin the subject catalyst composition comprised at least 99.7 weightpercent ²³⁸ U isotope.

EXAMPLE II

Catalyst A, prepared in accordance with Example I, was used in theproduction of methane-rich gas after being activated by being partiallyreduced in the following manner. Catalyst A was inserted in a fixed bedpilot plant reactor simulating the conditions of an adiabatic reactor ofindustrial size operated without recycle. The reactor was constructedfrom 54 inches of 1 inch, Sch. 80 316 stainless steel (SS) pipe and had316 SS seal-pieces at each bed against 4 inch diameter heads. The topseal-piece was provided with a central inlet whereas the bottomseal-piece had a side outlet. The reaction chamber nominally contained100 cc of Catalyst A in a 10-inch bed.

Catalyst A was treated within the reactor with 2,000 standard cubiccentimeters per minute of pure hydrogen for a period of 10 hours at atemperature of 850° F. and a pressure of 200 psig. This was followed byan additional 6 hours of treatment with hydrogen at the same flow andtemperature conditions but at a pressure of 320 psig. By means of thepreviously described treatment with hydrogen, Catalyst A was partiallyreduced so that about 50-60 weight percent of the nickel was present inthe metallic state. The uranium was present in the activated catalystprimarily in the form of NiUO₄ and less than 5 weight percent of atleast one oxide of uranium selected from the group consisting of UO₃ andUO₂. The weight percent of uranium present in the activated catalyst,based on the total weight of the activated catalyst, was 64.2. Theweight ratio U/Ni in the activated catalyst was about 3.0.

Runs 1 to 6 in Table I below, illustrate the use of the activatedcatalyst of the subject invention as prepared in Examples I and II. Datais provided in Table I for the production of methane-rich gas comprisingat least 26 mole percent (dry basis) of CH₄ by passing a feed gascomprising CO and at least one member of the group consisting of H₂ andH₂ O through a fixed-bed reactor containing the subject catalyst. Meanswere employed in the process for cooling the off gas from the reactor tobelow the dew point of a normally liquid materials in the off gas, andcollecting and separating the liquids from the gas stream. Conventionalequipment for measuring and controlling the rate of flow of the feed andproduct streams, temperature and pressure, and sampling and meteringwere provided.

The methane content of the product gas, and accordingly the heatingvalue, was maximized i.e. 53.3-97.4 mole percent CH₄ in Runs 1-3 byintroducing a dry feed gas into the reaction zone. Further, in Runs 1-3there were about 100% conversions of CO and H₂. While the mole ratio H₂/CO was greater in Run 1 than in Run 3, the reaction conditions weremilder and the methane content of the product gas was higher in Run 1 incomparison with Run 3. Run 1 clearly shows the high activity of thesubject catalyst at low temperatures i.e. 395° F. by converting 100% ofthe CO and H₂ in the feed gas into a product gas containing 97.4 molepercent methane. A large amount of CO₂ was produced in the product gasin Run 2 by reducing the mole ratio H₂ /CO of the feed gas to thecatalytic reactor to about 1.

Run 4 shows that the subject catalyst has the property of catalyzing thewater-gas shift reaction. The feed gas in Run 4 comprises CO and H₂ O,but no hydrogen. The mole ratio H₂ O/CO was 1. Substantially all of theCO was converted in the reactor. First, by the water-gas shift reaction,a portion of the CO in the feed gas was reacted with H₂ O to produce H₂and CO₂. Then, the H₂ was reacted with the remainder of the CO toproduce CH₄ and CO₂. In Run 5, the feed gas contained H₂, CO, and H₂ O.The increased amount of H₂ in the feed resulted in an increased yield ofCH₄ in Run 5 in comparison with that produced in Run 4.

EXAMPLE III

Runs 5, 6, and 7 illustrate the thermal stability of the activatedcatalyst prepared according to this invention in comparison with acommercial nickel catalyst. In Runs 5 and 6, the average bedtemperatures for the subject catalyst ranged from about 904° to 1061°F., while maximum bed temperatures may range up to about 1350° to 1500°F. CO and H₂ conversions were high i.e. 64 to 100 percent; and, the dryproduct gas contained from about 33 to 50 mole percent CH₄. Forcomparison, Run 7 and made using a commercial nickel methanationcatalyst 104. However, in Run 7 rapid deactivation of the commercialnickel catalyst took place when the average temperature of the bedreached above 990° F. The percent conversion of CO and H₂ fell rapidlyto respectively 16 and 24.8. The mole percent methane in the dry productgas was only about 7.

EXAMPLE IV

The high percent conversion of CO and H₂ and the high yield of CH₄, asin Runs 1-3, illustrate the high activity of the catalyst of the subjectinvention in comparison with the typical commercial iron oxide FischerTropsch catalyst that was employed in Run 8. Further, carbon containingliquids were produced in Run 8, while none were produced in Runs 1-3.These differences clearly distinguish between the two catalysts. Thecatalysts of the subject invention produce mainly CH₄ at conditionswhere normal Fischer Tropsch catalysts produce multicarbon compounds.The comparatively high level of CO conversion at low H₂ /CO ratios alsopoints out the ability of the catalysts of the subject invention tocatalyze the water-gas shift reaction simultaneously with themethanation reaction.

The overall results as shown in Table I serve to point out that theactivated catalyst of the subject invention is very active, has goodselectivity to methane, and has good thermal resistance.

                                      TABLE I    __________________________________________________________________________                                                        Carbon    Feed Gas     Reaction Conditions          Product Gas                                                        Containing    Run Mole Ratio                 Pressure                      Temp. SV    Time                                      % Conversion                                              Mole % (dry Basis)                                                        Liquids    No. H.sub.2 /CO            H.sub.2 O/CO                 Atm. °F.                            vol/vol/hr                                  Sec.                                      CO  H.sub.2                                              CO.sub.2                                                   CH.sub.4                                                        %     Catalyst    __________________________________________________________________________    1   3.13            0    100  395   2494  6.5 100 100 2.6  97.4 0     (a)    2   1.06            0    100  709   1226  9.6 100 100 46.7 53.3 0     (a)    3   1.70            0    300  583   1633  22.3                                      100 99.6                                              5.2  93.0 0     (a)    4   0   1    300  877   2477  6   99.9                                          --  73.4 26.0 0     (a)    5   1.00            0.80 300  1061  4917  3.7 99.7                                          100 49.6 50.4 0     (a)    6   3.19            0    300  904.6 7602  3.7 100 64  11.1 33   0     (a)    7   2.24            0    300  900   4561  6.3 16  24.8                                              1.1  7.0  0     (b)    8   1.85            0    300  490     582 68.6                                      43.7                                          25.5                                              --   6.3  6.2   (c)    __________________________________________________________________________     (a) Catalyst A prepared in Example I and activated in Example II.     (b) Commercial nickel methanation catalyst 104.     (c) Commercial iron oxide Fischer Tropsch catalyst.

It will be evident to those skilled in the art that variousmodifications of this invention can be made, or followed, in the lightof the foregoing disclosure and discussion, without departing from thespirit or scope thereof.

We claim:
 1. A process for the production of a methane-containing gascomprising:(1) introducing a gas mixture comprising CO and at least onemember of the group consisting of H₂ and H₂ O at a temperature in therange of about 375° to 1500° F. and a pressure in the range of about 1to 300 atmospheres into a reaction zone containing an unsupportedcatalyst comprising an alkali-metal promoted partially reduced mixtureof at least one nickel uranate and at least one oxide of nickel, whereinthe total weight percent of uranium present in said catalyst basis totalweight of said catalyst composition is in the range of over 50 to about90 weight percent; (2) converting CO in said gas mixture into CH₄ byreacting CO in said reaction zone with hydrogen which is either (a)introduced with said gas mixture, (b) produced by the water-gas shiftreaction in said reaction zone, or both (a) and (b); and (3) removing aCH₄ -containing gas from said reaction zone.
 2. The process of claim 1wherein the mole ratio of either H₂ /CO or H₂ O/CO in the feed gasmixture reacted in (2) is in the range of 0 to about 10 and the othermole ratio is in the range of about 0.3 to
 10. 3. The process of claim 1wherein the minimum atomic ratio H/C of the gas mixture reacted in (2)is
 2. 4. The process of claim 1 provided with the steps of cooling theCH₄ -containing gas stream from (2) below the dew point, and separatingwater to produce a dry gas stream comprising at least 25 mole % CH₄. 5.The process of claim 4 provided with the step of contacting said dry gasstream with a solvent absorbent and removing CO₂ therefrom.
 6. Theprocess of claim 1 wherein the weight ratio U/Ni in the catalyst in (1)is in the range of about 7-1, and at least 30 to about 80 weight percentof the nickel in the catalyst is in the metallic state.
 7. The processof claim 1 where in step (1) said nickel uranate is at least one memberselected from the group consisting of NiUO₄ and NiU₃ O₁₀, and said oxideof nickel is at least one member selected from the group consisting ofNiO and Ni₃ O₄.
 8. The process of claim 1 where in step (1) saidalkali-metal promotor is at least one member selected from the groupconsisting of potassium, sodium, and cesium.
 9. The process of claim 1wherein the uranium in the catalyst composition in (1) comprises atleast 99.7 weight percent ²³⁸ U isotope.
 10. A process for theproduction of a methane-containing gas comprising:(1) reacting ahydrocarbonaceous material by partial oxidation with a free-oxygencontaining gas optionally in the presence of a temperature moderator atan autogenous temperature in the range of about 1700° to 3000° F. and apressure in the range of about 1 to 300 atmospheres in the reaction zoneof a free-flow unpacked non catalytic synthesis gas generator to producean effluent gas stream comprising H₂, CO, H₂ O, and at least one memberselected from the group consisting of CO₂, CH₄, COS, H₂ S, Ar, N₂,particulate carbon, and ash; (2) cooling, cleaning and optionallypurifying at least a portion of the gas stream from (1) to obtain a feedgas stream comprising carbon monoxide and H₂ and/or H₂ O and at leastone member selected from the group consisting of CH₄, CO₂, N₂, and Ar;(3) introducing the cooled, cleaned and purified feed gas stream from(2) into a reaction zone containing an unsupported catalyst consistingof an alkali-metal promoted partially reduced mixture of at least onenickel uranate selected from the group consisting of NiUO₄ and NiU₃ O₁₀and at least one oxide of nickel selected from the group consisting ofNiO and Ni₃ O₄ ; and wherein the total weight of uranium present in thecatalyst in (3) based on the total weight of said catalyst compositionis in the range of over 50 to about 90 weight percent, the weight ratioU/Ni in the catalyst in (3) is in the range of about 5-2, and at least30 to about 80 weight percent of the nickel in the catalyst in (3) is inthe metallic state; and reacting said feed gas stream at a temperaturein the range of about 375° to 1500° F. and at a pressure in the range ofabout 1 to 300 atmospheres; and (4) removing a methane-containingproduct gas stream from said reaction zone in (3).
 11. The process ofclaim 10 wherein the mole ratio of either H₂ /CO or H₂ O/CO in the feedgas mixture reacted in (3) is in the range of 0 to about 5 and the othermole ratio is in the range of about 0.3 to
 5. 12. The process of claim10 wherein the uranium in the catalyst composition in step (3) comprisesat least 99.7 wt. % ²³⁸ U isotope.
 13. The process of claim 10 whereinthe feed gas stream from (2) is introduced into a fixed bed of catalystin step (3) at a space velocity in the range of 350-10,000vol./vol./hr..
 14. The process of claim 10 wherein the alkali-metalpromoter in step (3) is at least one member selected from the groupconsisting of potassium, sodium, and cesium.
 15. The process of claim 10wherein the catalyst composition in step (3) before partial reductioncomprises a mixture of the following materials in weight percent: NiO1.0 to 30, Ni₃ O₄ trace to 30, NiUO₄ 70 to 98, and K₂ O 0.01 to 2.0. 16.The process of claim 10 wherein the catalyst composition in step (3)before partial reduction comprises a mixture of the following materialsin weight percent: NiO 1.0 to 30, NiU₃ O₁₀ 70 to 98, and K₂ O 0.01 to2.0.
 17. The process of claim 10 wherein the catalyst composition instep (3) before partial reduction comprises a mixture of the followingmaterials in weight percent: NiO 1.0 to 30, NiUO₄ 30 to 68, and NiU₃ O₁₀30 to 68, and K₂ O 0.01 to 2.0.
 18. The process of claim 10 wherein thepressure in steps (2) through (4) is substantially the same as thepressure in the gas generator in step (1) less ordinary drop in thelines.
 19. The process of claim 10 provided with the step of cooling thegas stream from (4) by indirect heat exchange with the purified feed gasstream from (2).
 20. The process of claim 10 provided with the addedstep of cooling all or a portion of the gas stream from (4) below thedew point and condensing out and separating H₂ O from said gas stream toproduce a CH₄ -rich gas stream comprising at least 25 volume percentCH₄.
 21. The process of claim 20 wherein 1-25 volumes of said dewateredgas per volume of fresh feed gas from step (2) are recycled through thecatalytic reaction zone in step (3).
 22. The process of claim 20provided with the added step of removing CO₂ from the dewatered CH₄-rich gas stream.
 23. The process of claim 22 with the added step ofrecycling the CO₂ separated from the process gas stream to the partialoxidation gas generator in (1) as at least a portion of the temperaturemoderator.
 24. The process of claim 10 wherein said hydrocarbonaceousfeed is selected from the group consisting of petroleum distillate andresidua, gas oil, residual fuel, reduced crude, whole crude, asphalt,coal tar, coal derived oil, shale oil, tar sand oil, and mixturesthereof.
 25. The process of claim 10 wherein said hydrocarbonaceous feedis a pumpable slurry of solid carbonaceous feed selected from the groupconsisting of coal, lignite, coal char, particulate carbon, petroleumcoke, and mixtures thereof in H₂ O or in a liquid hydrocarbonaceousfuel.
 26. The process of claim 10 wherein said hydrocarbonaceous feed isa liquid hydrocarbon material containing combined oxygen selected fromthe group consisting of alcohols, ketones, aldehydes, organic acids,esters, ethers, and mixtures thereof, optionally in admixture with aliquid petroleum material.
 27. The process of claim 10 wherein saidfree-oxygen containing gas is selected from the group consisting of air,oxygen-enriched air (22 mole percent O₂ and higher), and preferablysubstantially pure oxygen (95 mole percent O₂ and higher.
 28. Theprocess of claim 10 wherein the catalyst in step (3) is promoted withabout 0.01 to 2.0 weight percent of an oxide of said alkali-metal. 29.The process of claim 10 wherein the unpurified portion of said cooledand cleaned gas stream from step (2) is mixed with the purified portionand the mixture is introduced into the reaction zone in step (3) with orwithout admixture with recycle gas.